Continuous process for alkylating an aromatic hydrocarbon



Nov. 28, 1967 N. MOULDEN 3,355,508

CONTINUOUS PROCESS FOR ALKYLATING AN AROMATIC HYDROCARBON Filed July 5, 1965 4 Sheets-Sheet 1 12 MAKE-UP PARAFFIN CHLORINATION 4 HC'I.

1 T K ATION 1 MAKE UP r-' 5 AL YL BENZENE RED on.

RED OIL ALUMINUM 2ND.ALKYLATION n------- Hcl BENZENE 151'. FRACTIONATION PARAFF'N 2ND FRACTIONATION RECYCLE Y LIGHT ALKYLATE 3RD. FRACTIONATION HEART CUT ALKYLATE 4TH. FRACTIONATION V HEAVY ALKYLATE |NVENTOR HOWARD N. MOUL DEN 1". 5c; arib ATTORNEYS Nov. 28, 1967 H. N. MOULDEN 3,355,508

CONTINUOUS PROCESS FOR ALKYLATING AN AROMATIC HYDROCARBON Filed July 3, 1963 4 Sheets-Sheet 2 U A 3 O m k n.

I! LU m J I l- J .3 U I REACTOR SETTLER SPENT CATALYST FIG.2

REACTOR INVENTOR HOWARD N. MOULDEN ATTORNEYS Nov. 28, 1967 H. N. MOULDEN 3,355,508

CONTINUOUS PROCESS FOR ALKYLATING AN AROMATIC HYDRQCARBON United States Patent Ofi ice Patented Nov. 28, 1967 3,355,508 CGNTINUQUS PRQCESS FOR ALKYLATING AN AROMATHI HYBROCARBON Howard N. Mouiden, San Anselmo, Calif assignor to Chevron Research Company, a corporation of Delaware Filed July 3, 1963, Ser. No. 292,578 14 Claims. (Cl. 260-671) The present invention relates to a process for the alkylation of an aromatic hydrocarbon. More particularly, the invention has to do with a continuous process for alkylating an aromatic hydrocarbon, for example benzene, with a C9-C13 alkyl chloride. The alkylaryl hydrocarbon is a valuable intermediate in the production of synthetic detergents, wetting agents or surface active agents by sulfonation and neutralization, for example with sodium hydroxide, to produce sodium alkylaryl sulfonate.

It is well known to prepare alkylaryl compounds of the type above described by reacting or condensing an arcmatic hydrocarbon, for example benzene, with an alkyl chloride, usually an alkyl chloride mixture, in the presence of aluminum chloride catalyst. However, the use of aluminum chloride requires that certain precautions in handling be observed. The catalyst must be maintained anhydrous and, in certain situations, free-flowing. In a continuous process, owing to deactivation of the catalyst, elaborate means must be provided to ensure a fresh supply of active catalyst. Previously, in order to obtain substantially complete or quantitative alkylation, long reaction times, high concentration of catalyst, and high temperatures were used. In a batch process, although replenishment of catalyst need not be made, good conversions of alkyl chloride and high yield of desired alkylated product were difiicult to obtain even at long residence times.

In accordance with the present invention, there is provided a process which employs a catalyst system which is easier to handle and requires less stringent moisturefree conditions, and which process results in substantially complete utilization or conversion of the alkyl chloride material in a relatively short reaction time.

The aromatic hydrocarbon to be condensed with the alkyl chloride may be mononuclear in structure, as exemplified by the benzene series, including benzene itself, and its lower alkyl derivatives, such as toluene and xylene; or it may be polynuclear, as exemplified by naphthalene.

The alkylating alkyl chloride component is usually obtained as a mixture by the chlorination of suitable hydrocarbon stock. Examples of suitable hydrocarbon stock for chlorination are petroleum distillate fractions or cuts of satisfactory boiling range, such as kerosene; paraflins which result from the hydrogenation of fatty acids; hydrogenated products of ethylene polymer having 9 to 18 carbon atoms, and products of the Fischer-Tropsch reaction, having 9 to 18 carbon atoms obtained by the hydrogenation of carbon monoxide in the presence of a metal catalyst of the iron group.

Chlorination of the above hydrocarbon stock is effected in conventional fashion, as shown, for example, by U.S. Patent No. 2,499,578. Conditions are such as to'minirnize the formation of polychlorides, but to favor the desired monochloride.

Generally, a suitable chlorinated product is one in which the organically bound chlorine content is far less than would obtain if chlorine were combined with all of the hydrocarbon molecules on the basis of one chlorine atom per molecule of hydrocarbon. Usually, the degree of chlorination is expressed in terms of mol percent chlorination of the hydrocarbon stock, and conventional chlorination processes are carried out to the extent of about 10 to 50% chlorination, and more often to chlorination.

A particularly useful alkyl chloride alkylating component is one in which the alkyl group is essentially of straight-chain structure, whereby upon alkylation a relatively straight-chain alkylate of the aromatic compound is produced. Suitable stock for chlorination is, therefore, a straight-chain or normal paraifin mixture, such as those isolated from petroleum distillates by means of molecular sieves or by means of urea adduction. Here, as above indicated, chlorination is carried out to the extent of preferably 15 to 25%. With such stock, as with other less linear paraflin stock, the chlorine atom will attach to the paraifin or hydrocarbon molecule at different sites, i.e. to terminal carbon atoms, thus producing the primary chloride; or to a carbon atom more near the middle of the straight chain to produce a secondary chloride or a tertiary chloride where branching is present. The conventional chlorination process will produce a mixture of nalkyl monoand polychlorides. The monochloride content will predominate in secondary chlorides with a minor but substantial proportion of primary chlorides; the proportion of primary chloride will usually range from about 8 to 15%, by weight, and of secondary chloride, from about to 92%.

Alkylaromatics in which the alkyl groups are essentially straight chains are convertible by sulfonation and neutralization into a product asserted to be superior in certain respects to alkylarylsulfonate detergents, which have a more complex alkyl group, as typified by the sodium polypropylene benzene sulfonates, and which presently form a large share of the synthetic detergent market.

Thus, it appears that regardless of the ever increasing content of pollutants generally finding their way into sewage systems, the sodium polypropylenebenzene sulfonate type of detergent is at least partly blamed for interfering with proper and efiicient processing of sewage. Because of their branched structure, these detergents are biologically hard, i.e., are resistant to bacterial attack and degradation into less complex, simple molecules. Some of these are asserted to destroy bacteria useful in the digestion of organic pollutants and to pass unchanged through the sewage plant, causing undesirable foaming on the way, and thence to find their way into surface, ,as well as underground, waters, thus contaminating them.

Therefore, it appears that regardless of the degree of culpability of the polypropyleneben'zene sulfonate type detergent, it is advantageous to produce a detergent alkylate which is more biodegradable or biologically soft, i.e. readily digested by bacteria and which appears in sewage waters in low concentrations, or not at all, in the form of its original molecule.

Essentially, the invention involves a continuous twostep liquid phase alkylation process wherein an aromatic hydrocarbon, for example benzene, is contacted With an alkyl chloride containing 9 to 18 carbon atoms in a first reaction zone, in the presence of a red oil alkylation catalyst complex produced in a second alkylation zone. An excess of aromatic compound ranging from 2 to 25 mols, preferably 5 to 15 mols, per mol of alkyl chloride is used. Temperature of reaction in this zone is such as to cause condensation of aromatic compound with the alkyl chloride, and temperatures of the order of to 185 F., preferably to F., are cited as examples. Reaction is continued for about 3 to 15 minutes. Under the specified conditions, not all of the alkyl chloride is converted, but a major proportion, up to about 95% and higher, but below 100%, of the alkyl chloride is converted. After reaction in the first alkylation zone, the reaction product mixture is then treated to remove spent red oil catalyst, as by settling. While the process is operative using a treatment by which all of the red oil is separated, in practice a small amount of red oil will remain dissolved or entrained in the hydrocarbon-containing phase to produce a substantially red oil-free mixture.

The substantially red oil-free reaction product mixture, comprising alkylaromatic hydrocarbon, unreacted alkyl chloride, hydrogen chloride, and aromatic hydrocarbon, is passed into a second alkylation zone and there contacted with aluminum metal catalyst at alkylation temperatures for a period of time sufficient to convert substantially all of the alkyl chloride. Reaction temperatures and residence times will lie Within the range of 100 F. to 185 F. and 0.3 to 15 minutes, respectively.

The etfiuent from the second reaction zone is then separated into a red oil alkylation catalyst complex phase and an organic phase comprising alkylaromatic hydrocarbons. The alkylation catalyst complex phase is sent to the first alkylation zone as the alkylation catalyst for the reaction in that zone. From the organic alkylaromatic phase, desired alkylaromatic product is recovered.

In a still more specific embodiment of the invention, substantially the whole of the reaction product mixture resulting from the chlorination of n-paraffin stock is continuously charged to a first alkylation zone. Aklyl chlorides are difficult to separate from unchlorinated parafiin, and the fact that the whole mixture can be used Without removing the unchlorinated paraffin is a distinct advantage in that a costly separation process is not required, and the unreacted paraflin can serve as desirable diluent.

In carrying out this embodiment of the invention, it is often advantageous to treat the paraflin stock to make it even more desirable. The normal parafin content can be increased by subjecting the paraflin stock to a separation process whereby the normal paraflins are selectively separated from isoparaflins, aromatics, and cyclic parafiins. Conventional methods for the selective separation of normal paraflins are the molecular sieve types. One such method involves the use of urea whereby a urea adduct of the normal parafiin is formed and then isolated from the unadsorbed isoparaffins, aromatics, and cyclic paraffins. In another example of a molecular sieve-type separation process synthetic crystalline zeolites, for example calcium alumino silicates, are used.

As already indicated, the chlorination product mixture charged to the first alkylation zone will ordinarily contain to 50 mol percent, preferably to 25 mol percent, alkyl chlordies and 50 to 90 mol percent, preferably 75 to 85 mol percent, unchlorinated parafiins. In addition, the alkyl chloride content will ordinarily consist predominantly of secondary n-alkyl chlorides, for example 85 to 92 mol percent, and a minor but substantial portion of primary n-alkyl chlorides, for example 8 to 15 mol percent.

Along with the chlorination product mixture, an excess of aromatic hydrocarbon, for example benzene, is charged to the first a kylation zone. The product desired is one in which the alkyl group from the alkyl chloride replaces but one hydrogen atom of the aromatic nucleus to produce, for example a m-onoalkylbenzene. To achieve this result an excess of aromatic hydrocarbon ranging from about 2 to 25 mols, preferably 5 to 15 mols of aromatic hydrocarbon for each mol of alkyl chloride in the feed is used. The excess aromatic hydrocarbon and unreacted paraflin are both readily separable from final alkylaromatic compound, and can both be re-used, the benzene recycled to the alkylation zone and the parafiin chlorinnated to produce more alkyl chloride.

While the alkylation of the present process is not unduly sensitive to the presence of moisture, it is nevertheless advantageous to use fairly dry feed containing not more than 0.01% water.

In the first alkylation zone the feed, comprising alkyl chlorides, parafiins, and aromatic hydrocarbons, is subjected to alkylation conditions in the presence of red oil catalyst produced in a second zone. Ninety to ninety-five percent of the alkyl chlorides are converted in this zone. As a result, substantially all of the secondary chloride and a substantial portion of the primary chloride will ordinarily be converted in the first zone. Reaction conditions fall Within the range already given, i.e., temperatures of 100 F. to 185 F., and residence times of 3 to 15 minutes. Since the over-all process is a liquid phase one, it can be conducted under substantially atmospheric pressures, it being advantageous to employ a slight superatmospheric pressure, for example up to p.s.i.a., to facilitate flow of materials.

The reaction product mixture of the first alkylation zone will comprise red oil, alkylaromatic hydrocarbons, alkyl chlorides, aromatic hydrocarbon, and hydrogen chloride resulting from the condensation of aromatic hydrocarbon and alkyl chloride.

Red oil is separated from the reaction product mixture and the remainder is then passed into the second alkylation zone. Since much hydrogen chloride is produced, in practice it is vented from the reaction product mixture during alkylation or from the feed prior to its entry into the second alkylation zone, thus producing a feed containing dissolved hydrogen chloride or one saturated with hydrogen chloride. Such a feed entering the second alkylation zone, containing about, by weight, 1 to 30 parts of hydrogen chloride per 1000 parts of feed, is suitable in the commercial operation of this process.

The second alkylation zone may comprise a vessel to which all of the aluminum metal, such as turnings, shot, pellets, or pieces, which will be used in the process has been charged, i.e. a tower or the like prepacked with aluminum metal. In this case, it will he found advantageous to control, as by venting or stripping, the quantity of hydrogen chloride entering the second reaction zone with red oil-free efiluent from the first reaction zone beyond the amount needed to promote reaction. Too much hydrogen chloride uses up aluminum metal unnecessarily, forms more red oil than is necessary, and thus results in Waste of metal and in lesser utilization of catalyst activity. In general, satisfactory operation ensues if the feed to the second alkylation zone contains dissolved hydrogen chloride to the extent above specified.

A convenient way of contacting aluminum metal and efiiuent from the first alkylation zone is to continuously introduce aluminum metal, for example in the form of pellets, into the second alkylation zone as needed, along with the feed to this zone, a satisfactory rate being 0.03 to 0.1 gram atom per gram atom of chlorine in the chlorinated paraffin feed charged to the first alkylation Zone. In this case, the aluminum metal itself is the limiting factor in its being used up and in the rate of catalyst formation, and therefore no great precaution need be taken to remove hydrogen chloride over and above any required to promote alkylation. To achieve substantially complete conversion of alkyl chloride, reaction temperatures of to F. at residence times of 0.3 to 15 minutes are used.

From the efliuent of the second reaction zone, a red oil catalyst phase is separated and is sent to the first alkylation zone as the catalyst complex used in that zone. The red oil-free product, comprising unreacted hydrocarbon, alkylaromatic hydrocarbon and parafiins is treated to recover desired alkylaromatic product, while the aromatic compound and parafiins may be reused in the process.

Further to illustrate the practice of the invention, reference is now made to the drawings, it being understood that conventional pieces of equipment, such as valves, heat exchangers, reflux drums, reboilers, evacuators, and the like are omitted for the sake of simplicity.

Although the chlorinated paraffin feed can be obtained from an independent off-site source, and its preparation need not be integrated with the alkylation and fractionation steps, FIG. 1 is a flow sheet illustrating in block diagram the major features and process steps involved in a preferred embodiment of the invention for the pr duction of detergent alkylate from chlorination to fractionation;

FIG. 2 is a diagrammatic elevation of apparatus for carrying out one embodiment of the invention wherein alkyl chloride is continuously contacted with aromatic compound in a first alkylation zone followed by contact of efiiuent from this reaction zone in a second alkylation zone with aluminum metal, red oil catalyst formed in this zone being sent to the first alkylation zone;

FIG. 3 shows diagrammatically a modified type of apparatus for carrying out another embodiment of the invention wherein chlorinated parafiins, together with unchlorinated paraflins, are used as feed for contact with the aromatic hydrocarbon in a first alkylation zone in the presence of red oil catalyst and excess aromatic hydrocarbon, and efiluent from this alkylation zone continuously contacted with aluminum metal in a second reaction zone, red oil formed in the second reaction zone being continuously sent to the first alkylation zone; and

FIG. 4 shows in diagrammatic form principal apparatus used in the chlorination step.

FIGS. 3 and 4 show, when placed side by side, a complete process for making detergent alkylate from chlorination to fractionation.

According to FIG. 1, chlorine and parafiin feed are subjected to chlorination conditions in a chlorination zone, more particularly illustrated by FIG. 4, to produce a 15 to 25% chlorination product. During the chlorination reaction, hydrogen chloride is formed, and it can be vented directly from the chlorination zone or from the chlorination product eflluent.

The chlorination product mixture comprising chlorinated paratfins and unchlorinated paraflins then flows into the first alkylation zone, together with more than enough aromatic hydrocarbon, for example benzene, to condense with the alkyl chloride, generally 5 to 15 mols benzene per mol of alkyl chloride. Alkylation is performed in the presence of red oil catalyst complex flowing into first alkylation zone from the second alkylation zone. Proportions of red oil catalyst depend on temperature and time. Thus, less red oil catalyst will require higher temperatures and/or longer reaction times. In general, the amount of red oil catalyst will range from about 5 to 25 volume percent based on total reaction mixture, including red oil catalyst.

Alkylation conditions already alluded to and more specifically mentioned in connection with FIGS. 2 and 3 can effect conversion of about 90 to 95% of the alkyl chloride in a period of 3 to 15 minutes. Excess hydrogen chloride can also be vented from the alkylation zone.

Spent red oil catalyst leaves the first alkylation zone in approximately an amount equal to that of fresh red oil catalyst entering the first alkylation zone from the second alkylation zone.

The reaction product mixture of the first alkylation zone comprising alkylaromatic compounds, unreacted aromatic compound, unreacted alkyl chloride, and hydrogen chloride then passes into the second alkylation zone, which may comprise a stirred tank reactor or a column packed with aluminum metal. In this zone it is contacted with red oil catalyst and/or aluminum metal which may be simultaneously introduced into the second alkylation zone.

Conditions of alkylation mentioned above and in connection with FIGS. 2 and 3 are such as to convert substantially all of the alkyl chloride entering into the second alkylation zone at a residence time ranging from 0.3 to 15 minutes.

Red oil complex formed in the second alkylation zone passes into the first alkylation zone as already indicated, while the red oil-free product comprising unreacted aromatic hydrocarbon, alkylaromatics, unreacted paraifins and hydrogen chloride is treated to remove acidic products and then conducted into the fractionation zone.

From this zone the benzene is recycled to first alkylation zone, together with required make-up benzene.

The reaction product mixture leaving the second alky1a tion zone is then treated to isolate the alkylaromatics, or detergent alkylate. This may be accomplished by using conventional means.

As above indicated, a preferred procedure involves the chlorination of a C C paraifin mixture to give a partially chlorinated mixture comprising chlorinated paraffins and a major proportion of unchlorinated parafiins. After alkylation there will, therefore, be produced a mixture comprising alkylaromatics and unchlorinated paraffins. Separation or isolation of the alkylaromatics can be accomplished, using conventional means. One such method involves the selective sulfonation of the aromatic nucleus of the alkylaromatics, leaving the paraflins unsulfonated. The unsulfonated paraffins can then be removed from the sulfonation mixture by solvent extraction, for example with pentane.

However, it will often be found advantageous to use fractional distillation. When this method is used, cleancut separations between alkylaromatics and paraffins, because of overlap in boiling ranges, are difficult to obtain where the whole C C paraffin feed has been used in the beginning. For example, C C n-parafiins boil within about the range 300 F. to 605 F.; and the corresponding C C alkylbenzenes, between about 520 F. to 730 F.

It will therefore be found advantageous to start with parafiin feed cut or fractions of narrower boiling ranges, i.e., parafiin cuts made up of no more than four molecular species, whereby a clean-cut separation of alkylaromatics and parafiins can be effected by distillation. Thus, where the feed is a C C fraction, the boiling range will be about 345 F. to 455 F.; for a C C fraction, 385 F. to 490 F.; and for a C C fraction, 420 F. to 520 F.

Conditions in the first fractionation, hereinabove and hereafter typified by benzene, Will be such as to separate the benzene. Accordingly, all material boiling below 250 F. to 275 F. will be removed.

After the first fractionation, parafiin is removed in a second fractionation zone and is recycled to the chlorination zone. Conditions will vary depending on boiling point or boiling point range of the paraflin stock fed to the process. Thus, material will be recycled to the chlorination zone which boils below 460 F. for a C -C n parafiin cut; below about 520 F. for a C C cut; and below about 490 F. for a G -C cut.

Light alkykate, that is, alkylate which will not convert into a satisfactory detergent, is next taken off. For a C -C this light alkylate will be material boiling below about 520 P.; for a C C cut, below about 580 F.; and for a C C cut, below about 545 F.

Finally, heart-cut alkylate is taken off as the desired fraction. For a C C cut, the desired heart cut will be material boiling below 625 F.; for a C C below 650 F.; and for a C -C below about 675 F., material boiling higher than 625 F., 650 F., and 675 F. going to heavy alkylate.

To summarize, for a C -C benzene alkylate suitable for conversion to a superior detergent, the boiling range will be between about 520 F. and 625 F.; for a C C alkylate, between 545 F. and 650 F.; and for a C -C alkylate, between about 580 F. and 675 F.

Referring now more specifically to FIG. 2, a chloroparaffin of the desired carbon content, together with an excess of aromatic hydrocarbon, using benzene as illustrative, required to react with all of the chloroparafi'ln, enters through line 1 into dehydrator 2, wherein residual moisture from the reactants is removed. The dehydrator is provided with a water adsorbent, for example silica gel or a molecular sieve-type water adsorbent. The dried mixture of reactants then passes from the dehydrator through line 3 into first alkylation reactor 4. Reactor 4 may be a vessel of the stirred reactor type. Thorough mixing is eifected as with stirrer impelled by motor 6. Reaction in reactor 4 is conducted in the presence of red oil catalyst complex obtained from the second alkylation reactor, flowing into the first alkylation reactor through line 7.

Reaction temperatures range from 100-185 F. during a residence time or 3 to 15 minutes. Under these conditions, 90 to 95% of the alkyl chloride is converted.

During reaction in the first reactor, continuous vent ing of excess free undissolved hydrogen chloride gas occurs through line 8. Line 8 passes through heat exchanger 9 to cool and condense light boiling materials, such as benzene, which returns to the reaction zone.

Reaction product mixture continuously leaves reactor 4 and through line 10 passes into a settling zone represented by settler 11. A lower catalyst phase separates, and part of it is discarded through line 14 in a quantity approximately equivalent to that coming from the second stage. The volume of red oil catalyst in reactor will be kept within the range of 5 to 25%, while the volume of red oil catalyst in settler 11 will be kept at a minimum consistent with good operation practice, so as to avoid excessive dilution of more reactive red oil catalyst from the second stage. The red oil not discarded is recycled back to reactor 4 through line 13.

The upper organic phase in settler 11, now comprising alkylaromatic hydrocarbon, excess benzene, and unreacted alkyl chloride, and containing dissolved hydrogen chloride, is continuously withdrawn from settler 11 through line 15 and flows into the second alkylation reactor 16. As indicated in the drawing, reactor 16 is a tower packed with aluminum metal, such as aluminum turnin'gs, bird shot, pellets, or lumps.

Conditions of reaction in reactor 16 are such as to effect substantially complete conversion of the alkyl chloride fed thereinto. Accordingly, temperatures of reaction will vary from 100-l85 F., and contact times from 0.3 to 15 minutes.

During reaction in reactor 16, additional hydrogen chloride and hydrogen are vented and passed through line 17, through heat exchanger 18, whereby desired volatiles, other than hydrogen and hydrogen chloride, are condensed and returned to the reactor.

The reaction product mixture of reactor 16 passes through line 19 into a second settler 20. In settler 29 a red oil catalyst separates and is passed through line 7 into reactor 4 as its catalytic agent. An upper organic phase comprising alkylbenzene hydrocarbons and excess benzene is withdrawn through line 21 and fractionated to recover desired alkylbenzene alkylate.

Reactor 16 instead of a packed tower can be a stirred tank reactor to which aluminum is fed at a predetermined rate. In this case, a major proportion of the red oil from settler 26 is recycled to the stirred reactor, while the remainder is sent to the first reactor.

Referring to FIG. 3, a feed representing substantially the whole reaction product mixture after partial chlorination of a parafiin cut of the desired boiling range is introduced through lines 59, 51 and 52 into first alkylation reactor 53, together with dry aromatic hydrocarbon, hereinbelow typified by benzene, through line 54. An excess of benzene ranging from about 5 to 15 mols per mol of alkyl chloride is cited as illustrative. Reaction conditions are such as to cause 90 to 95% conversion of the alkyl chloride by maintaining a temperature Within the range 100-185 F., and a contact time of 3 to 15 minutes, under substantially atmospheric pressure. Catalyst from a second alkylation zone, shown in greater detail below, flows simultaneously into reactor 53 through line 55 in an amount such that it will form 5 to volume percent of the total reaction product mixture in the reactor.

Thorough mixing of reactants and catalyst is desired. This may be accomplished by l'me mixers in the feed lines and by type of reactor used. Accordingly, reactor 53 is represented as being a so-called continuous loop reactor in which its contents continuously circulate from bottom to top through an external loop represented by lines 52 and 56.

Alkylation product mixture continuously leaves reactor 53 through line 57 and passes into separator or settler 58. Undissolved hydrogen chloride gas and benzene are discharged through vent line 59 into benzene absorber or gas scrubber 60. A red oil phase separates out in separator 58 and leaves via line 61. A portion of this red oil returns through line 62 to reactor 53, and a minor portion corresponding in quantity to red oil coming from the second alkylation zone, more fully described below, is discharged through line 63.

The upper organic phase of settler 58 leaves through line 64 and passes through line 65 into a second alkylation reactor 66. Reactor 66 is of a type similar to reactor 53. T 0 provide for thorough agitation, the contents of reactor 66 continuously leave at the bottom through line 67, pass through line 65, and return to the top. Aluminum metal is continuously added through line 68 at a rate based on the amount of chlorine contained in the mixture introduced through line 50, a satisfactory amount being 0.03 to 0.1 gram atomper gram atom of chlorine. As in the first reactor, the amount of red oil is maintained within the range of 5 to 25 volume percent of the total mixture in reactor 66.

Reaction conditions in reactor 66 are such as to cause substantially complete conversion of the alkyl chloride content introduced thereinto. Temperatures are of the order of 100l85 F., and residence times, of the order of 0.3 to 15 minutes.

Reaction product mixture of reactor 66 leaves through line 69 and passes into an expansion chamber '71 wherein disengagement of gas from the liquid reaction mixture occurs. Gas is removed through line 71 and passed into absorber-condenser 72 where it is washed with water introduced through line 73, volatiles not in solution being vented through line 74.

In absorber-condenser 72, hydrogen chloride and water form an aqueous solution. This solution passes through line 75 into separator 76, wherein an aqueous hydrogen chloride phase and a benzene phase are formed. The aqueous phase leaves through line '77. The benzene phase leaves through line 78 and joins line 83 to form line 84.

The degassed mixture is passed from expansion chamber '71) through line 79 into separator or settler 80. In separator 80 a lower red oil phase and an upper organic phase are formed. The red oil phase is withdrawn through line 81. Part of the red oil is recycled to reactor 66 through lines 82 and 65, and part, corresponding substantially to that made from the aluminum metal, is sent to reactor 53 through line 55.

The upper organic phase of separator 80 is withdrawn through line 83, joins line 78 to form line 84, and then combines with heavy alkylate flowing through line 158, FIG. 4. The resulting solution goes through line 85 and is thoroughly mixed with caustic introduced through line 86. The resulting mixture is passed through line 87 in separator or settler 88. To effect thorough mixing, the lines may be provided with line mixers, not shown. In vessel 88, neutralization of residual acidic contaminants is effected and then separated. Spent caustic leaves through valved line 89, and a portion of it is recycled through line 9% into line 86, the remainder going to disposal through line 91. Fresh or make-up caustic is introduced through line 92 in an amount equal to that going to disposal.

The caustic-treated product in vessel 88 is withdrawn through line 93, mixed with water introduced through line 94, and sent through line 95 into washing vessel and separator 96. Fresh water is introduced through line 97 and waste water removed through valved line 98 and sent to disposal through line 99. A portion of the water, if desired, may be recycled through line 108 to line 94.

The washed product from vessel 96 is then removed through line 101 and sent to settler or coalescer 102. Water is separated in settler 102 and sent to disposal through line 103.

Substantially water-free product is passed from settler 102 through line 104 into distillation column 105. Benzene is distilled overhead, is condensed by means not shown and taken through lines 106 and 107 into dehydrator 108. Prior to dehydration, make-up benzene is introduced through line 109. In order to facilitate continuous operation, two or more of these dehydrators may be connected in parallel so that one may be regenerated while the other is on stream. Dry benzene is then passed through lines 54, 51 and 52 into reactor 53.

The bottom fraction of distillation column 105 is withdrawn through line 110 and is passed into another distillation column 111. Unchlorinated parafi'ins are fractionally separated and withdrawn through line 112 and sent with make-up paraffin from line 113, if desired, through line 114 to chlorination, FIG. 4 below. The bottom fraction of distillation column 111 passes through line 115 and enters a third distillation column 116.

Low boiling hydrocarbons are withdrawn from distillation column 116 through line 117, While the higher boiling fraction is withdrawn through line 118 and introduced into another distillation column 119. The overhead fraction containing the desired product is removed through line 120 and a bottom heavy alkylate is withdrawn through line 121, a portion being passed through line 157, FIG. 4.

For eflicient operation, vent line 59 carrying benzene and hydrogen chloride empties into gas scrubber 60, wherein the benzene is recovered and recycled through lines 122, 51, and 52 into reactor 53. Scrubbing is efiected by countercurrent treatment with n-paraflin feed stock, charged through line 123.

In an integrated, efliciently run operation, the apparatus and facilities for chlorination of n-parafiins will tie in with apparatus and facilities for the alkylation reactions described above. In such an integrated process, makeup parafiin, i.e., an amount of parafiin corresponding to that used in the chlorination operation can be introduced directly into the chlorinator, as for example through line 113.

However, the present process afiords an opportunity to condition make-up paraflin and hence all of the parafiin used after start up, to minimize undesirable side reactions. Paratiin feed, after being subjected to a molecular sieve treatment but prior to chlorination, will contain undesirable constituents which are not removed. Such impurities can be removed from parafiin to be chlorinated, if the make-up paraffin is first passed through the alkylation operation prior to being sent to chlorination. In addition, the make-up paraflin can aid in the scrubbing operation in scrubber 60. Accordingly, make-up paraffin is passed through line 124 through dehydrator 125 for removal of any moisture, and then through line 123- into scrubber 60. Any gases in scrubber 60 are vented through line 126. From the scrubber the parafiin is sent with the recovered benzene through lines 122, 51, and 52 into reactor 53.

Referring now to FIG. 4, n-parafiin from distillation column 111, FIG. 3, flows through lines 112 and 114, FIGS. 3 and 4, into a first chlorinator 135. Chlorine gas passes through valved line 136 and thence into line 137 into chlorinator 135.

The chlorine gas and n-parafiin are subjected to conditions of chlorination such as to give a maximum amount of monochloride and a minimum amount of dichloride. Accordingly, the chlorination reaction is carried out at a temperature in the range 75-250 F., at a pressure ranging from atmospheric to 65 p.s.i.a. Low chlorination temperatures, for example 100-l50 F., preferably 110-130 F., can be used by activating the chlorination reaction with light, and higher temperatures, e.g. 150-200 F. thermally,

10 preferably 160-180" F. The amount of chlorine gas passed into the chlorinator is such as to give 1015% chlorina tion; that is, the amount of chlorine gas entering the chlorinator will be in the range of 0.1 to 0.15 mol per mol of entering parafiin.

In the first chlorinator, chlorination proceeds to about 40-60% of total paraffin conversion desired. The partially chlorinated reaction product mixture then leaves through line 138 and enters into separator 139, wherein hydrogen chloride gas and unreacted chlorine are separated from the partially chlorinated reaction product mixture. The hydrogen chloride gas and excess chlorine leave the separator through line 140, and the degassed reaction product mixture leaves the separator through line 141 and flows into second chlorinator 142.

Chlorination to produce a product of 20-30% total chlorination is effected by contacting the partially chlorinated mixture with additional chlorine gas, 0 .1 to 0.15 mol per mol of parafiin, entering the second chlorinator from valved line 136 through line 143. Conditions for completing the chlorination in chlorinator 142 are substantially similar to those employed in the first chlorinator, namely, temperatures in the range 75-250 F. and pressures of up to 65 p.s.i.a.; as in the first chlorinator, lower temperatures of reaction can be used when using light as a reaction activator.

Upon completion of the chlorination reaction, the chlorination product leaves the second chlorinator through line 144 and flows into second separator 145, wherein excess chlorine gas and hydrogen chloride are again freed from the reaction product mitxure, and leave the separator through line 146 and enter line 140.

The liquid product of separator 145 leaves through line 148 and enters chlorine stripper 149, wherein residual chlorine gas is stripped from the product mixture. To expedite the stripping action, a gas inert to the system, such as nitrogen, is passed through the mixture. Conveniently, the stripping gas is dry hydrogen chloride obtained from gas scrubber 60, FIG. 3, passing into the stripper through line 150.

The stripped chloroparaffins leave chlorine stripper 149 through line 151 and flow, if desired, into reactor 53, FIG. 3, through line 50. The stripped gas and some nparafiin leave stripper 149 through line 152, enter line 140, and thence pass into separator 147. The liquid phase of separator 147 comprising the n-paraflin can be sent back to chlorinator 142 through line 153. The overhead from separator 147, comprising chlorine, hydrogen chloride, and some n-paraflin, leaves through line 154 and passes into paraffin absorber, scrubber, or stripper 155. In the parafiin absorber, chlorine and hydrogen chloride leave through vent line 156. To facilitate the separation of chlorine and hydrogen chloride gas from the n-paraflins, heavy alkylate is introduced into paraffin absorber 155 through line 157. The heavy alkylate used is conveniently obtained from column 119 via line 121, FIG. 3. The heavy alkylate and n-parafiin leave parafin absorber through line 158 to join line 84, and thence into neutralization and separation vessel 88, FIG. 3.

The practice of the invention will be further illustrated by the following examples.

Example 1 A mixture of n-parafiins boiling over a range of 340 F. to 490 F. had the following composition on a mol basis: C1, C11, C12, C C14, 17.8%; C 0.5%; C 0.1%; remainder, impurities. The n-parafiinic hydrocarbon mixture was partially chlorinated in the liquid phase at -130 F. to give a chlorination reaction product mixture of alkyl chlorides, 88 mol percent secondary and 12 mol percent primary, the total chlorine content by weight being 3.54%. The chlorination product mixture, pounds (21.6 gallons), consisting essentially of 16 mol percent of alkyl chlorides and 1 1 84 mol percent of uuchlorinated parafiinic hydrocarbons was mixed with 110 pounds (15 gallons) of benzene.

The benzene-parafiin-alkyl chloride mixture was charged at a rate of 14 gallons per hour into the first stage rector, containing 3 gallons of reaction mixture.

The volume of the reaction mixture was held a constant 3 gallons by continuously withdrawing it. In the first-stage reactor, temperature was maintained at 163 F., and the pressure was maintained at 35 p.s.i.a. The reaction mixture was mixed throughout the run.

Reaction product was withdrawn and was fed into the first settler, in which the red oil catalyst phase was separated from the hydrocarbon phase and was recycled back into the first reactor at a rate such as to maintain a constant red oil volume of 510%. A portion of the red oil phase in the first settler, amounting to about 0.4 pound per hour, was removed and discarded. The amount of spent red oil discarded was such as to maintain a constant first stage inventory of red oil.

The organic phase from the first settler was then passed into the second reactor. Analysis by means of gas-liquid chromatography showed that 95% of the alkyl chloride had been converted in the first reactor. The second reactor contained 3 gallons of reaction mixture and was maintained at a constant volume by continuously withdrawing reaction mixture.

Aluminum shot, about %-inch in diameter, was fed to the second reactor at a rate of 0.053 pound per hour. Addition of the aluminum was carried out under a nitrogen blanket in order to prevent any moisture from entering the reactor. The second reactor was maintained at 160 F. at a pressure of 38 p.s.i.a. The reaction mixture was mixed throughout the run.

The reaction product from the second reactor was passed into a second settler, in which the dense red oil catalyst phase was separated from the hydrocarbon phase. Red oil catalyst was recycled to the second reactor, so that its red oil content was 20 volume percent. Red oil, in excess of a predetermined second stage inventory, was fed to the first reactor.

The hydrocarbon phase from the second settler was removed continuously to maintain a constant volume in the settler. This withdrawn material was passed into a caustic scrubber wherein residual acidic compounds were neutralized, thence into a water wash wherein all salts and residual caustic were removed, and finally into barrels for storage.

At steady state, analysis of a sample of the reaction product by means of gas-liquid chromatography indicated a 99.8% conversion. After the removal of the benzene, the reaction product pounds) gave, on distillation, the following fractions:

Parafiin cut: B.P. 350491 R; 8.11 lbs.

Light alkylate cut: Bl 49l548 F.; 0.04 lb. Detergent alkylate cut: 3.1 548639 F.; 1.67 lb. Bottoms: B.P. 639 F.; 0.18 lb.

The yield of detergent alkylate was 93 weight percent, based on alkyl chlorides, or 105 weight percent (88 mol percent), based on monochloride.

Example 2 The feed stock for this example was 10 mols of henzene and 4 mols of straight-chain parafiin per mol of chloroparafiin. The paraffin and chloroparaffin had the following mass mol distribution: C 1%; C 22.2%; C12, C13, C14, 17.4%; and C15, 0.7%.

Apparatus similar to that of FIG. 2 was used. The first reactor was maintained at 165 F. with a red oil catalyst phase of 9.8 volume percent. The average residence time was set at 11.8 minutes. All of the red oil formed in the second stage was pumped to the first stage.

Under these conditions, the total conversion of alkyl chlorides in the first reactor was 96.4 mol percent.

T he product from the first stage was continuously passed into a settler. The red oil phase from the settler was recycled into the first reactor except for a small amount removed to maintain a constant red oil volume in the first reactor of 9.8 volume percent. The hydrocarbon phase saturated with hydrogen chloride at 165 F. and one atmosphere of pressure was passed into an aluminum packed tower, one-inch in diameter, and eight inches long, and containing 135 grams of aluminum rivets.

The reactants passed into the tower at a liquid hourly space velocity of 10.1 v./v./hr. in an essentially plug fiow. The average residence time in this column was 3 minutes at a temperature of 165 F. Under these conditions, no alkyl chloride could be detected in the product. At the end of the run the aluminum shot was weighed and the quantity used during the run was calculated to be 0.06 gram atom of aluminum/mol of alkyl chloride introduced into the first zone. Over weight percent yield of alkylbenzene boiling within the range 544635 F. was realized upon distillation of the product.

Example 3 A hydrocarbon feed stock obtained from distillation of crude oil and having a boiling point spread of from 325 to 475 F. is treated with molecular sieves to obtain an essentially pure n-paraffin fraction. The process is conducted substantially according to the procedure indicated in FIG. 3, parts being by weight.

The n-parafiin material is fed to the first alkylation reactor via line 124. Recycle n-paraflins from the alkylation zone, containing once-through make-up n-parafiins, are fed into the chlorinator through line 114 at a rate of 1000 parts/ hour. Simultaneously, chlorine gas containing 2.5 weight percent hydrogen chloride is fed at a rate of 59 parts/ hour through line 136 into the same chlorinator 135. Prior to entering the chlorinator the chlorine gas passes through a Pyrex pipe illuminated with a fluorescent light.

The chlorination vessel, one in which efiicient stirring and heat removal can be effected, having a volume of reaction mixture such as to provide a residence time of 15 minutes, is maintained at 120 F. and 35 p.s.i.a.

The reaction mixture is continuously removed through line 138 into the separator 139.

Hydrogen chloride and chlorine are separated and removed through line 140 at a rate of 15 and 2 parts/ hour, respectively. This gas stream also removes a small amount of paraffin vapor. The degassed liquid is then passed into the second chlorinator 142 through line 141. Simultaneously, chlorine gas of the same quantity and quality as before passes into the second chlorinator through a fiuorescent lighted Pyrex pipe in line 143.

The second chlorinator is of the same type and size as the first. It operates under the same conditions. The reaction mixture is withdrawn continuously through line 144 and sent into second separator 145. Hydrogen chloride and chlorine are removed through line 146 at 30 and 2 parts/hour, respectively. At the same time, a small amount of paraiin vapor is removed in the vent gas.

The liquid bottoms material from this separator is withdrawn through line 148 at a rate of 1068 parts/hour, heated to F., and fed to the chlorine stripper 149. Here the material is sparged with 27 parts/ hour of hydrogen chloride obtained from the alkylation section and introduced through line 150. This gas contains about 1%, by weight, of n-parafiins. The volatile materials pass out of the stripper via line 152 into the general gas line 144] at a rate of 31 parts/hour of hydrogen chloride, 7 parts/ hour of chlorine, and 1 part/hour of parafiins.

The ofigases from lines 140, 146, and 152 are combined, cooled to 100 F., and passed into separator 147. Here 1 part/hour of condensed parafiin is removed and returned to the second chlorinator via line 153. The exit gases are passed through line 154 into a ring packed parafiin scrubber or stripper '155. In this scrubber, the

gases pass upward through a downward flow of 42 parts/ hour of heavy alkylate introduced through line 157 at 100 F. The ofigas, 88 parts/hour, is treated to recover hydrogen chloride and chlorine. The heavy alkylate containing about 1.5%, by Weight, n-paraflins passes out of the parafiin scrubber at 43 parts/hour through line 158 to the caustic neutralizer.

The liquid product from the chlorine scrubber 149 consisting of, by weight, 75.9% n-parafiins, 23.4% chlorinated parafiins, and 0.7% hydrogen chloride leaves through line 151 and is passed at 1056 parts/hour via line 50 into the alkylation reactor 53.

Simultaneously, 1096 parts/hour of dried benzene containing 9.4%, by weight, make-up benzene and 0.7% of dissolved n-paraflin is passed through line 54 into reactor 53. Make-up paraffin is fed at 204 parts/hour into the benzene absorber 60 in which 23 parts/hour of benzene are absorbed from the ottgas stream 59 and the resulting solution is fed into reactor 53 via line 122.

A portion of the red oil catalyst from the second settler 80 leaves through line 81 and is fed via line 55 at 35 parts/hour into the first reactor. Agitation is maintained in the first reactor 53 by continuously pumping the liquid contents out of the bottom of the reactor and back into the top of the reactor via lines 56 and '52. The reactor contents are maintained at 160 F. and 35 p.s.i.a.

The reaction mixture is Withdrawn from reactor 53 at 2766 parts/hour through line 57 in order to maintain a liquid residence time of 15 minutes. The withdrawn material is passed to settler 58. Here the red oil phase separates and is withdrawn through line 61 and recycled back to the reactor through line 62 at 350 parts/hour. A portion of the red oil, 23 parts/hour, is discarded through line 63.

Hydrogen chloride, 36 parts/hour, and benzene, 23 parts/hour, are removed from the settler 58 through line 59 and passed through benzene absorber 60. In this absorber, the benzene is separated from the hydrogen chloride by scrubbing with 204 parts/hour of fresh n-paraffins fed through line 123. Prior to use in the benzene absorber, the n-paraffins are dried by passing through desiccant packed dehydrator 125. The n-paratfins-benzene mixture with about 1% absorbed hydrogen chloride is passed into reactor 53 through line 122.

The red oil-free material from settler 58 is passed at 2334 parts/hour into the second reactor 66 via lines 64 and 65. At the same time aluminum shot is fed to this reactor via line 68 at 1 part/hour.

In this reactor, reaction conditions are 160 F. and 20 p.s.i.a. Agitation is etfected by continuously pumping reaction product out of the bottom of the reactor back into the top of the reactor via lines 67 and 65. This reactor is comparable in size to the first reactor, and a 15-minute residence time is maintained by withdrawing reaction mixture at 2685 parts/hour through line 69. This reaction mixture passes via line 69 into an expansion chamber 70 wherein 13 parts/hour hydrogen chloride and 21 parts/hour of benzene are. removed through line 71 and. fed to the absorber-condenser 72. The hydrogen formed in this reactor is removed through, lines 71 and 74.

In vessel 72 the vent gases are contacted with 26 parts/ hour of water introduced through line 73 to. form 33 weight percent hydrochloric acid. This aqueous-hydrocarbon mixture is then removed through line 75 and passed into separator 76... In this vessel the phases are separated and. the aqueous acid is removed through line 77 at 39 parts/hour. The hydrocarbon phase, 21 parts/ hour, is pumped via line 78 into the overflow line 83 from settler 80.

The hydrocarbon material from expansion chamber 70, 2651 parts/hour, is cooled to 125 F. and passed to settler 80 via line 79,. The dense red oil phase is. removed through line 81 at 385 parts/hour. Of this red oil, 350 parts/hour is recycled back to reactor 66 via line 82 and 14 35 parts/hour is fed to the first reactor through line 55 as fresh make-up catalyst.

The hydrocarbon phase of settler is withdrawn via line 83 at 2266 parts/hour and combined with 21 parts/ hour of benzene, line 78, from the separator 76 and 43 parts/hour of heavy alkylate, line 158, .from the chlorine stripper 155. All of this material, along with 53 parts/ hour of fresh 25 Baum caustic, is fed to the caustic neutralizer '88. To provide good mixing, caustic is recycled through lines 89, 90, and 86, except for an amount sufficient to balance the make-up caustic. This spent caustic is removed through line 91.

The neutralized hydrocarbon is next passed into a Water Washer 96 via lines 93 and 95. Fresh water is also added through line 97. To provide good mixing, water is recycled through lines 98, 10,0, and 94, except for a portion sufficient to balance the fresh incoming water. This spent water is discarded through line 99.

Water-washed hydrocarbon passes from vessel 96 into a coalescer 102 via line 101. Here the aqueous phase is removed via line 103 and the remainder, 2321 parts/ hour, is heated and passed to the benzene still 105 via line 104.

The benzene still, 105, is a 30-sieve tray continuous distillation column operating at an overhead temperature of 224 F. and 30 p.s.i.a. Benzene, 994 parts/hour and containing 0.7%, by weight, n-parafiins, is removed via line 106 and recycled back to the reactor after mixing with 102 parts/hour of make-up benzene introduced through line 109.

The bottoms from this still, 1327 parts/hour, are then passed to the parafiin still 111 via line 110. The paraffin still is a 40-sieve tray continuous distillation column operating with an overhead temperature of 325 F. at 4 p.s.i.a. The paraffin material taken overhead, 1000 parts/hour, is removed via line 112 and recycled to the first chlorinator 135.

The bottoms of the parafiin still are withdrawn through line 115 and pumped into the light alkylate still 116 at 327 parts/hour. This is a 30-sieve tray continuous distillation column operating at an overhead temperature of 410 F. at 4 p.s.i.a. The overhead material is removed through line 117 at 11 parts/hour and sent to storage.

The bottoms of the light alkylate still are removed through line 118 and pumped to the rerun still 119 at 316 parts/hour. The rerun still is a 30sieve tray continuous distillation column operating at an overhead temperature of 442 F. at 2 p.s.i.a. The material taken overhead is alkylbenzene product and is withdrawn through line 120 at a rate of 244 parts/hour. The bottoms from this still are taken to storage via line 121 at a rate of 72 parts/hour, except for 42 parts/hour which is sent to the paraffin scrubber via line 157.

I claim:

1. Continuous process for the production of detergent alkylate by the alkylation of an aromatic hydrocarbon with alkyl chloride in a first alkylation zone followed by alkylation in a second alkylation zone, which comprises subjecting a stoichiometric excess of an aromatic hydrocarbon compound and a mixture of C -C alkyl chlorides consisting predominantly of secondary alkyl chlorides and a minor but substantial portion of primary alkyl chlorides to alkylating conditions in a first alkylation zone in the presence of a red oil catalyst complex obtained from the second alkylation zone, said red oil catalyst complex constituting about 5 to 25 percent by volume of the contents of said first alkylation zone, continuing alkylation in the first alkylation zone until about 90-95% of the alkyl chlorides have been converted, passing the efiluent from the first alkylation zone to a first settling zone to form a first red oil phase and a first organic phase comprising alkylaromatic hydrocarbons, alkyl chlorides, aromatic hydrocarbon compound, and hydrogen chloride, recycling a portion of the red oil catalyst phase to the first alkylation zone, discharging a portion of the red 15 oil phase corresponding in amount to that supplied from the second alkylation zone, so as to maintain a volume of to 25 percent by volume of red oil catalyst complex in the first alkylation zone, passing the first organic phase into a second alkylation zone, and therein contacting it with aluminum metal under alkylation conditions, continuing the alkylation until substantially all of the alkyl chloride has been converted, thereby producing a second reaction product mixture comprising alkylaromatic hydrocarbons, aromatic hydrocarbons, and red oil catalyst complex, passing the efiiluent from the second alkylation zone to a second settler to form a second organic phase and a second red oil catalyst phase, passing said red oil catalyst to the first alkylation zone, and recovering alkylaromatic hydrocarbons from the second organic phase.

2. Process according to claim 1, wherein the red oil catalyst is present in the second alkylation zone in an amount of 5 to 25 percent by volume based on the contents of said second alkylation zone.

3. Process according to claim 2, wherein aluminum metal is continuously fed into the second alkylation zone at a rate of 0.02 to 0.1 gram atom per gram atom of chlorine in the feed entering the first alkylation zone.

4. Process according to claim 3, wherein the aromatic hydrocarbon is benzene and the alkylation reaction in the first and second alkylation zones is carried out at a temperature in the range of 100 to 185 F.

5. Process according to claim 4, wherein the first organic phase entering the second alkylation zone is saturated with hydrogen chloride.

6. Continuous process for producing a biologically soft detergent alkylate, which comprises reacting normal parafiinic hydrocarbons containing 9 to 18 carbon atoms per molecule with elemental chlorine in a chlorination zone to produce a chlorination product mixture consisting essentially of 10 to 50 mol percent chlorinated paraffins and 50 to 90 mol percent of unchlorinated paratfins, said chlorinated paraffins consisting predominantly of secondary parafiin chlorides and a minor but substantial portion of primary parafiin chlorides passing the chlorination product, and a substantial molar excess relative to the chlorinated parafiins of an aromatic hydrocarbon into a first alkylation zone and therein contacting them with a red oil alkylation catalyst complex at a temperature in the range of 100-185 F. for a period of 3 to minutes, to convert 90-95% of the chlorinated parafiins and to form a reaction product mixture comprising alkylaromatics, hydrogen chloride, unreacted aromatic hydrocarbon, and chlorinated paratiins, passing said last-mentioned alkylaromatics, unreacted hydrocarbon, chlorinated paraflins, and hydrogen chloride to initiate alkylation, into a second reaction zone and therein contacting them with aluminum metal at a temperature of 100-185 F. until substantially all of the chlorinated parafiins have reacted, separating from the efiiuent from the second reaction zone a red oil alkylation catalyst complex phase and an organic phase comprising alkylaromatics, l1n reacted aromatic hydrocarbons, and unchlorinated paraffins, passing the red oil alkylation catalyst complex phase into the first reaction zone to supply substantially all of the red oil alkylation catalyst complex used in that zone, and recovering biologically soft detergent alkylate from the organic phase. 1

7. Process according to claim 6, wherein the aromatic hydrocarbon is benzene and is present in proportions of 5 to 15 mols per mol of chlorinated paraffins.

8. Continuous process for producing a biologically soft detergent alkylate, which comprises reacting in a chlorination zone a mixture of normal paraffinic hydrocarbons of 9 to 18 carbon atoms, said mixture containing not more than four adjacent molecular species, with elemental chlorine to produce a chlorination product mixture consisting essentially of 10 to 50 mol percent chlorinated parafiins and 50 to 90 mol percent of unchlorinated parafiins, passing the chlorination product mixture and a substantial molar excess relative to the chlorinated parafiins of benzene into a first alkylation zone and therein contacting them with a red oil alkylation catalyst complex at a temperature in the range 100 F. to 185 F. until at least about but less than all, of the chlorinated paraffins have been converted, there-by forming a first alkylation product mixture comprising alkylaromatics, hydrogen chloride, parafiinic hydrocarbons, unreacted benzene and chlorinated paraffius, passing the effluent of said first alkylation product mixture into a first settling zone to form a first upper organic phase and a first lower red oil phase, passing the first upper organic phase and hydrogen chloride to a second alkylation zone, and contacting it with aluminum metal at a temperature in the range F. to F. until substantially all of the chlorinated paraifins have been converted, passing the efiluent from the second alkylation zone into a second settling zone to form a second upper organic phase and a lower red oil catalyst phase, passing said red oil catalyst phase to the first alkylation zone, discarding red oil catalyst from the first settler in an amount corresponding to that supplied to the first alkylation zone from the second alkylation zone, and recycling the remainder red oil catalyst from the first settler to the first alkylation zone, fractionally distilling the second upper organic phase, to separate benzene, detergent alkylate, and paraflinic hydrocarbons.

9. Process according to claim 8, wherein the parafiinic hydrocarbons boil in the range 300 Fsto 605 F., and the detergent alkylate, 520 F. to 730 F.

10. Process according to claim 8, wherein the benzene is recycled to the first alkylation zone, and the parafiinic hydrocarbons, together with a quantity of added normal pararfinic hydrocarbons approximately equal on a molar basis to the chlorinated paraflins, are sent to the chlorination zone.

11. Process according to claim 8, wherein the red oil catalyst in the first alkylation zone and in the second alkylation zone constitutes about 5 to 25 volume percent, based on the contents of said zones.

12. Process according to claim 9, wherein the paraffinic hydrocarbon feed stock to the chlorination zone has a boiling range of about 345 F. to 455 F., and the detergent alkylate, 520 F. to 625 F.

13. Process according to claim 9, wherein the paraifinic hydrocarbon feed stock to the chlorination zone has a boiling range of about 385 F. to 490 F., and the detergent alkylate, 545 F. and 650 F.

14. Process according to claim 9, wherein the paraflinic hydrocarbon feed stock to the chlorination zone has a boiling range of about 420 F. to 520 F., and the detergent alkylate, 580 F. to 675 F.

References Cited UNITED STATES PATENTS I 2,740,807 4/ 1956 Rappen et al. 260-671 3,078,322 2/1963 McCaulay 260683.5l 1,995,827 3/1935 Thomas 260-671 2,057,306 10/1936 Martin et al 252442 X 2,078,238 4/ 1937 Dreisbach.

2,233,408 3/1941 Flett 260-671 2,244,512 6/1941 Brandt 260671 2,388,428 11/1945 Mavity 260-671 3,169,987 2/ 1965 Bloch 260-67 1 OTHER REFERENCES Concise Chemical and Technical Dictionary, Second Edition, 1962, Chemical Publishing Co., Inc., New York, page 792.

The Petroleum Dictionary, University of Oklahoma Press, 1952, page 246.

DELBERT E. GANTZ, Primary Examiner. C. R. IQAXIS, Assistant Examiner. 

1. CONTINUOUS PROCESS FOR THE PRODUCTION OF DETERGENT ALKYLATE BY THE ALKYLATION OF AN AROMATIC HYDROCARBON WITH ALKYL CHLORIDE IN A FIRST ALKYLATION ZONE FOLLOWED BY ALKYLATION IN A SECONDALKYLATION ZONE, WHICH COMPRISES SUBJECTING A STOICHIOMETRIC EXCESS OF AN AROMATIC HYDROCARBON COMPOUND AND A MIXTURE OF C9-C18 ALKYL CHLORIDES CONSISTING PREDOMINANTLY OF SECONDARY ALKYL CHLORIDES AND A MINOR BUT SUBSTANTIAL PORTION OF PRIMARY ALKYL CHLORIDES TO ALKYLATING CONDITIONS IN A FIRST ALKYLATION ZONE IN THE PRESENCE OF A RED OIL CATALYST COMPLEX OBTAINED FROM THE SECOND ALKYLATION ZONE, SAID RED OIL CATALYST COMPLEX CONSTITUTING ABOUT 5 TO 25 PERCENT BY VOLUME OF THE CONTENTS OF SAID FIRST ALKYLATION ZONE, CONTINUING ALKYLATIN IN THE FIRST ALKYLATION ZONE UNTIL ABOUT 90-95% OF THE ALKYL CHLORIDES HAVE BEEN OCNVERTED, PASSING THE EFFLUENT FROM THE FIRST ALKYLATION ZONE TO A FIRST SETTLING ZONE TO FORM A FIRST RED OIL PHASE AND A FIRST ORGANIC PHASE COMPRISING ALKYLAROMATIC HYDROCARBONS, ALKYL CHLORIDES, AROMATIC HYDROCARBON COMPOUND, AND HYDROGEN CHLORIDE, RECYCLING A PORTION OF THE RED OIL CATALYST PHASE TO THE FIRST ALKYLATION ZONE, DISCHARGING A PORTION OF THE RED OIL PHASE CORRESPONDING IN AMOUNT TO THAT SUPPLIED FROM THE SECOND ALKYLATION ZONE, SO AS TO MAINTAIN A VOLUME OF 5 TO 25 PERCENT BY VOLUME OF RED OIL CATALYST COMPLEX IN THE FIRST ALKYLATION ZONE, PASSING THE FIRST ORGANIC PHASE INTO A SECOND ALKYLATION ZONE, AND THEREIN CONTACTING IT WITH ALUMINUM METAL UNDER ALKYLATION CONDITIONS, CONTINUING THE ALKYLATION UNTIL SUBSTANTIALLY ALL OF THE ALKYL CHLORIDE HAS BEEN CONVERTED, THEREBY PRODUCING A SECOND REACTION PRODUCT MIXTURE COMPRISING ALKYLAROMATIC HYDROCARBONS, AROMATIC HYDROCARBONS, AND RED OIL CATALYST COMPLEX, PASSING THE EFFLUENT FROM THE SECOND ALKYLATION ZONE TO A SECONMD SETTLER TO FORM A SECOND ORGANIC PHASE AND A SECOND RED OIL CATALYST PHASE, PASSING SAID RED OIL CATALYST TO THE FIRST ALKYLATION ZONE, AND RECOVERING ALKYLAROMATIC HYDROCARBONS FORM THE SECOND ORGANIC PHASE 